The present invention relates to a process for producing polypropylene from C3 olefins selectively produced from a catalytically cracked or thermally cracked naphtha stream.
The need for low emissions fuels has created an increased demand for light olefins used in alkylation, oligomerization, MTBE and ETBE synthesis processes. In addition, a low-cost supply of light olefins, particularly propylene, continues to be in demand to serve as feedstock for polyolefin, particularly polypropylene.
Fixed bed processes for light paraffin dehydrogenation have recently attracted renewed interest for increasing olefin production. However, these types of processes typically require relatively large capital investments and high operating costs. It is therefore advantageous to increase olefin yield using processes, which require relatively small capital investment. It would be particularly advantageous to increase olefin yield in catalytic cracking processes so that the olefins could be further processed into polymers such as polypropylene.
A problem inherent in producing olefins products using FCC units is that the process depends on a specific catalyst balance to maximize production of light olefins while also achieving high conversion of the 650xc2x0 F.+(xcx9c340xc2x0 C.) feed components. In addition, even if a specific catalyst balance can be maintained to maximize overall olefin production, olefin selectivity is generally low because of undesirable side reactions, such as extensive cracking, isomerization, aromatization and hydrogen transfer reactions. Light saturated gases produced from undesirable side reactions result in increased costs to recover the desirable light olefins. Therefore, it is desirable to maximize olefin production in a process that allows a high degree of control over the selectivity of C3 and C4 olefins.
One embodiment of the present invention is a process for producing polypropylene comprising the steps of (a) feeding a naphtha stream comprising from about 10 to 30 wt. % paraffins and between about 15 to 70 wt. % olefins to a process unit comprising a reaction zone, a stripping zone, a catalyst regeneration zone, and a fractionation zone; (b) contacting the naphtha stream with a fluidized bed of catalyst in the reaction zone to form a cracked product, the catalyst comprising a zeolite having an average pore diameter of less than about 0.7 nm and wherein the reaction zone is operated at a temperature from about 500xc2x0 to 650xc2x0 C., a hydrocarbon partial pressure of 10 to 40 psia (about 70-about 280 kPa), a hydrocarbon residence time of 1 to 10 seconds, and a catalyst to feed weight ratio between about 4 and about 10, thereby producing a reaction product wherein no more than about 20 wt. % of paraffins are converted to olefins and wherein propylene comprises at least about 90 mol. % of the total C3 products; (c) passing the catalyst through said stripping zone; (d) passing the stripped catalyst from the stripping zone to the catalyst regeneration zone where the catalyst is regenerated in the presence of an oxygen-containing gas; (e) recycling the regenerated catalyst to the reaction zone; (f) fractionating the cracked product to produce a C3 fraction, a C4 fraction rich in olefins, and optionally a C5 fraction rich in olefins; (g) passing the C4 fraction to the reaction zone or the stripping zone, or both; and, (h) separating propylene from the C3 fraction and polymerizing the propylene to form polypropylene.
In another embodiment of the present invention the catalyst is a ZSM-5 type catalyst.
In an embodiment of the present invention a C5 fraction rich in olefins is also recycled.
In another embodiment of the present invention the feedstock contains about 10 to 30 wt. % paraffins, and from about 20 to 70 wt. % olefins.
In another embodiment of the present invention the reaction zone is operated at a temperature from about 525xc2x0 C. to about 600xc2x0 C.
Feedstreams that are suitable for producing the relatively high C2, C3, and C4 olefin yields are those streams boiling in the naphtha range containing from about 5 wt. % to about 35 wt. %, preferably from about 10 wt. % to about 30 wt. %, and more preferably from about 10 to 25 wt. % paraffins, and from about 15 wt. %, preferably from about 20 wt. % to about 70 wt. % olefins. The feed may also contain naphthenes and aromatics. Naphtha boiling range streams are typically those having a boiling range from about 65xc2x0 F. to about 430xc2x0 F. (about 18xc2x0 C. to about 225xc2x0 C.), preferably from about 65xc2x0 F. to about 300xc2x0 F. (about 18xc2x0 C. to about 150xc2x0 C.).
The naphtha can be a thermally cracked or a catalytically-cracked naphtha. The naphtha streams can be derived from the fluid catalytic cracking (FCC) of gas oils and resids, or they can be derived from delayed or fluid coking of resids. Preferably, the naphtha streams used in the practice of the present invention derive from the fluid catalytic cracking of gas oils and resids. FCC naphthas are typically rich in olefins and/or diolefins and relatively lean in paraffins. It is within the scope of the instant invention to feed or co-feed other olefinic streams that are not catalytically- or thermally-cracked naphthas, such as an MTBE raffinate, into said reaction zone with the primary feed. It is believed that this will increase the yield of propylene.
The process of the present invention is performed in a process unit comprising a reaction zone, a stripping zone, a catalyst regeneration zone, and a fractionation zone. The naphtha feed is fed into the reaction zone where it contacts a source of hot, regenerated catalyst. The hot catalyst vaporizes and cracks the feed at a temperature from about 500xc2x0 C. to about 650xc2x0 C., preferably from about 525xc2x0 C. to about 600xc2x0 C. The cracking reaction deposits coke on the catalyst, thereby deactivating the catalyst. The cracked products are separated from the coked catalyst and sent to a fractionator. The coked catalyst passes through the stripping zone where a stripping medium, such as steam, strips volatiles from the catalyst particles. The stripping can be preformed under low-severity conditions to retain a greater fraction of adsorbed hydrocarbons for heat balance. The stripped catalyst is then passed to the regeneration zone where it is regenerated by burning coke on the catalyst in the presence of an oxygen containing gas, preferably air. Decoking restores catalyst activity and simultaneously heats the catalyst to a temperature from about 650xc2x0 C. to about 750xc2x0 C. The hot regenerated catalyst is then recycled to the reaction zone to react with fresh naphtha feed. Flue gas formed by burning coke in the regenerator may be treated for removal of particulates and for conversion of carbon monoxide. The cracked products from the reaction zone are sent to a fractionation zone where various products are recovered, particularly a C3 fraction, a C4 fraction, and optionally a C5 fraction. The C4 fraction and the C5 fraction will typically be rich in olefins. One or both of these fractions can be recycled to the reactor. They can be recycled to either the main section of the reactor, or a riser section, or a stripping section. It is preferred that they be recycled to the upper part of the stripping section, or stripping zone. Recycling one or both of these fractions will convert at least a portion of these olefins to propylene.
While attempts have been made to increase light olefins yields in the FCC process unit itself, the present invention uses its own distinct process unit, as previously described, which receives naphtha from a suitable source in the refinery. The reaction zone is operated at process conditions that will maximize C2 to C4 olefins (particularly propylene) selectivity with relatively high conversion of C5+ olefins. Suitable catalysts used with the present invention contain a crystalline zeolite having an average pore diameter less than about 0.7 nanometers (nm), said crystalline zeolite comprising from about 10 wt. % to about 50 wt. % of the total fluidized catalyst composition. It is preferred that the crystalline zeolite be selected from the family of medium-pore size ( less than 0.7 nm) crystalline aluminosilicates, otherwise referred to as zeolites. Of particular interest are the medium-pore zeolites with a silica to alumina molar ratio of less than about 75:1, preferably less than about 50:1, and more preferably less than about 40:1, although some embodiments may incorporate a silica to alumina ratio greater than 40:1. The pore diameter, also referred to as effective pore diameter, is measured using standard adsorption techniques and hydrocarbonaceous compounds of known minimum kinetic diameters. See Breck, Zeolite Molecular Sieves, 1974 and Anderson et al., J. Catalysis 58, 114 (1979), both of which are incorporated herein by reference.
Medium-pore size zeolites that can be used in the practice of the present invention are described in xe2x80x9cAtlas of Zeolite Structure Types,xe2x80x9d eds. W. H. Meier and D. H. Olson, Butterworth-Heineman, Third Edition, 1992, which is hereby incorporated by reference. The medium-pore size zeolites generally have a pore size from about 5 xc3x85, to about 7 xc3x85 and include for example, MFI, MFS, MEL, MTW, EUO, MTT, HEU, FER, and TON structure type zeolites (IUPAC Commission of Zeolite Nomenclature). Non-limiting examples of such medium-pore size zeolites, include ZSM-5, ZSM-12, ZSM-22, ZSM-23, ZSM-34, ZSM-35, ZSM-38, ZSM-48, ZSM-50, silicalite, and silicalite 2. The most preferred is ZSM-5, which is described in U.S. Pat. Nos. 3,702,886 and 3,770,614. ZSM-11 is described in U.S. Pat. No. 3,709,979; ZSM-12 in U.S. Pat. No. 3,832,449; ZSM-21 and ZSM-38 in U.S. Pat. No. 3,948,758; ZSM-23 in U.S. Pat. No. 4,076,842; and ZSM-35 in U.S. Pat. No. 4,016,245. All of the above patents are incorporated herein by reference. Other suitable medium-pore size zeolites include the silicoaluminophosphates (SAPO), such as SAPO-4 and SAPO-11 which is described in U.S. Pat. No. 4,440,871; chromosilicates; gallium silicates; iron silicates; aluminum phosphates (ALPO), such as ALPO-11 described in U.S. Pat. No. 4,310,440; titanium aluminosilicates (TASO), such as TASO-45 described in EP-A No. 229,295; boron silicates, described in U.S. Pat. No. 4,254,297; titanium aluminophosphates (TAPO), such as TAPO-11 described in U.S. Pat. No. 4,500,651; and iron aluminosilicates.
The medium-pore-size zeolites can include xe2x80x9ccrystalline admixturesxe2x80x9d which are thought to be the result of faults occurring within the crystal or crystalline area during the synthesis of the zeolites. Examples of crystalline admixtures of ZSM-5 and ZSM-11 are disclosed in U.S. Pat. No. 4,229,424, which is incorporated herein by reference. The crystalline admixtures are themselves medium-pore-size zeolites and are not to be confused with physical admixtures of zeolites in which distinct crystals of crystallites of different zeolites are physically present in the same catalyst composite or hydrothermal reaction mixtures.
The catalysts of the present invention may be held together with an inorganic oxide matrix material component. The inorganic oxide matrix component binds the catalyst components together so that the catalyst product is hard enough to survive interparticle and reactor wall collisions. The inorganic oxide matrix can be made from an inorganic oxide sol or gel which is dried to xe2x80x9cbindxe2x80x9d the catalyst components together. Preferably, the inorganic oxide matrix is not catalytically active and will be comprised of oxides of silicon and aluminum. It is also preferred that separate alumina phases be incorporated into the inorganic oxide matrix. Species of aluminum oxyhydroxides-xcex3-alumina, boehmite, diaspore, and transitional aluminas such as xcex1-alumina, xcex2-alumina, xcex3-alumina, xcex4-alumina, xcex5-alumina, xcexa-alumina, and xcfx81-alumina can be employed. Preferably, the alumina species is an aluminum trihydroxide such as gibbsite, bayerite, nordstrandite, or doyelite. The matrix material may also contain phosphorous or aluminum phosphate.
Process conditions include temperatures from about 500xc2x0 C. to about 650xc2x0 C., preferably from about 500xc2x0 C. to 600xc2x0 C.; hydrocarbon partial pressures from about 10 to 40 psia (about 70-about 280 kPa) to about, preferably from about 20 to 35 psia (about 140-about 245 kPa); and a catalyst to naphtha (wt/wt) ratio from about 3 to 12, preferably from about 4 to 10, where catalyst weight is total weight of the catalyst composite. Steam may be concurrently introduced with the naphtha stream into the reaction zone, with the steam comprising up to about 50 wt. % of the naphtha feed. Preferably, the naphtha residence time in the reaction zone is less than about 10 seconds, for example from about 1 to 10 seconds. The reaction conditions will be such that at least about 60 wt. % of the C5+ olefins in the naphtha stream are converted to C4xe2x88x92 products and less than about 25 wt. %, preferably less than about 20 wt. % of the paraffins are converted to C4xe2x88x92 products, and that propylene comprises at least about 90 mol. %, preferably greater than about 95 mol. % of the total C3 reaction products with the weight ratio of propylene/total C2xe2x88x92 products greater than about 3.5.
Preferably, ethylene comprises at least about 90 mol. % of the C2 products, with the weight ratio of propylene:ethylene being greater than about 4, and that the xe2x80x9cfull rangexe2x80x9d C5+ naphtha product is enhanced in both motor and research octanes relative to the naphtha feed. It is within the scope of this invention to pre-coke the catalysts before introducing the feed to further improve the selectivity to propylene. It is also within the scope of this invention to feed an effective amount of single-ring aromatics to the reaction zone to also improve the selectivity of propylene versus ethylene. The aromatics may be from an external source such as a reforming process unit or they may consist of heavy naphtha recycle product from the instant process.